Polymers of vinyl monomers usually referred to as vinyl polymers and the polymeric reaction of these monomers has become an extremely important part in the syntheses of polymers. Such as polystyrene, polyvinyl chloride, poly(meth) acrylates, polyvinyl acetate, ABS resins and all kinds being widely used plastics, rubber, thermoplastic elastomers, coatings, etc. are related with such monomers. The most of them are obtained by free radical polymerization and almost occupy more than half in the polymers. However, due to the fast polymerization speed and large heat release of these polymers and therefore, rapid increases of the melt viscosity, the withdrawal of reaction heat and the diffusion and uniform distribution of the monomer become very difficult to synthesize by the bulk polymerization. Therefore, the solution, suspension, emulsion and other polymerization methods used to carry out instead of bulk polymerization.
Although the solution polymerization method can avoid the difficulties of the bulk polymerization, this polymerization method is exclusively used for theoretical studies because of the slow polymerization rate and the high energy consumption and environmental pollution caused by removing the solvent during the post-treatment process. Therefore, this polymerization method has never been industrialized, exclusive the anionic polymerization, which has to polymerize by this method. Suspension polymerization also avoids the difficulties of heat dissipation and high viscosity. Furthermore, it can easily separate the polymer from water and need not remove the solvent as usually in the solution polymerization process, therefore, it avoids unnecessary energy consumption. For these reasons, it is a successful industrial polymerization method. However, during the process of separation and purification of the polymerized crude product, serious environmental pollution may be caused by having to wash off the large amount of suspending agent and surfactant existing on the surface of the bead polymer. Furthermore, drying the crude product also consumes considerable energy. Besides, due to the residual surfactant and suspending agent, the purity and mechanical properties of the product will be lost for a certain extent. Therefore, this method of polymerization is only used in expandable polystyrene (EPS) processes where polymer properties do not be required very high and the crude polymer beads can be used directly, so that it is very economic. Emulsion polymerization is similar to suspension polymerization. However, because the molecular weight by it can be quite high at higher polymerization rates, which is more favorable for improving the product performance. On the other hand, the problem of environmental pollution is even worse because of the smaller polymer beads that make their surface emulsifiers and surfactants more difficult to remove. Fortunately, the polymers obtained by emulsion polymerization can be directly used as coating, with the result to reduce the energy consumption. However, the process to prepare plastics has brought great pollution, such as emulsion polymerization for ABS production.
The most currently popular polymerization method is the continuous bulk thermal polymerization. This polymerization method has no other impurities entering the system, and polymer is obtained only directly from the monomer, so the product can be of high transparency and environmental pollution also be dropped to a low level. Therefore, the method is very popular in the international. However, this polymerization method has a palpable defect that it is difficult to obtain high the molecular weight at the higher polymerization rate, therefore, the production efficiency drops substantially with increasing molecular weight. For example, if the weight average molecular weight of polystyrene is increased from 270000˜280000 to 400000, production efficiency will drop to 40% of normal conditions. Furthermore, in the latter part of the reaction, because of too large a viscosity, the system lacks mass transfer, insulting a very wide molecular weight distribution. Obviously, there always have been some defects in polymerization technique how to carry out the free radical polymerization. Precisely because of this, the research on the polymerization of vinyl monomers by reactive extrusion has emerged one after another.
Stuber et al. used a φ34 mm reverse self-cleaning tight-intermeshing twin-screw extruder to study the bulk polymerization of methyl methacrylate and determined the residence time and distribution of the material by injecting a solid dye into the first barrel. (Stuber N P, Tirrell M. Polym Process Eng, 1985, 3:71).
Lee (Lee R W, Miloscia W J. Standard Oil, U.S. Pat. No. 4,410,659, 1983) and Bodolus et al. (Bodolus C L, Woodhead D A, Standard Oil, U.S. Pat. No. 4,542,189, 1985) used a co-rotating twin screw extruder to research the bulk polymerization of methyl methacrylate with acrylonitrile and nitrile rubber dissolved in monomer. Polymerization conditions: input liquid monomer per minute 27 g (75 parts of acrylonitrile in it), 2.45 g of nitrile rubber, barrel temperature of 110˜177° C., screw speed of 75 rpm, 4 minutes of the residence time of material in the extruder, unreacted monomer was withdrawn from the devolatilizer on the extruder. The product had a 77% yield and the impact resistance of it was more than 10 times higher than that without the nitrile rubber.
With regard to the copolymerization of styrene with its corresponding monomer, Stober and Amos (U.S. Pat. No. 2,530,409, 1950) reported their research early in 1950, in which the styrene prepolymer was fed into a single screw extruder with an average residence time of 18 h, a temperature gradient along the screw of 120-200° C. and a screw speed of 1 rpm.
Illing (U.S. Pat. No. 3,536,680, 1970) investigated the bulk copolymerization of styrene with acrylonitrile, methyl methacrylate or acrylamide. In order to provide longer residence time, a 3-order tandem tightly intermeshing twin-screw extrusion reactor was designed. A 5° C. mixture of initiator-containing styrene and acrylonitrile monomer was fed from the feed inlet on the first twin-screw extruder, the material was heated in the extruder to 130-180° C. for 20-40 seconds then fed to second twin screw extruder of a about 60-200 mm in diameter. The monomers are mainly polymerized in the second screw extruder and the reactants flow in a thin layer in order to obtain sufficient and effective kneading. 1.52×105 Pa pressure was kept in the barrel, and the reaction residence time controlled in the range of 1.5 to 18 minutes. The reactant was then fed to third twin-screw extruder where the material was devolatilized and the unreacted monomer was removed from two evacuating ports.
However, as indicated in the above continuous bulk thermal polymerization method, in the polymerization of the vinyl monomer initiated by free radical, there is always a contradiction between the rate of polymerization and the molecular weight of the polymer. In order to achieve a sufficiently high molecular weight, the polymerization time must be very long, however, this is unacceptable for the reactive extrusion process based on continuously rapid production. Conversely, if the polymerization rate needs to be increased, the molecular weight and the conversion rate will be very low. Thus, radical-initiating reactive extrusion polymerization has so far been seldom truly commercialized.
CN 1587292B, CN 1587288A, CN1644597 have published a technique using the anionic polymerization to synthesize some plastics, rubbers and thermoplastic elastomers through the reactive extrusion bulk-polymerization in order to make the bulk reactive extrusion polymerization of the vinyl monomers into industrialization. CN101824151B has published a synthesis technology of fluorosilicone through the reactive extrusion of anionic ring-opening polymerization. However, due to the type of vinyl monomer being suitable for the anionic polymerization is too less, furthermore, the monomers and all the raw materials need to be refined extremely. Although anionic ring-opening polymerization can expand some application area of reactive extrusive polymerization, the channels of reactive extrusion polymerization initiated by the radical must be opened up, since the radical polymerization has the terrifically broad application.
As mentioned above, the early radical-initiated reactive extrusion polymerization is difficult to be practiced due to the larger technical defects. Thus, Chen Jixin et al. (Chemical Science and Technology, 2012, 20(2): 80-84) published a technique, in which MMA monomer pre-polymerizes at first in a polymerization kettle, then enters the screw extruder for the final polymerization with high viscosity. This proposal seems reasonable, but it is virtually impossible to implement. Since there exists a very high risk of burst polymerization while MMA monomer pre-polymerize in the kettle, in which a higher viscosity polymerization system cannot be rapidly cooled and terminated. Even if fortunately the burst polymerization did not happen, it is impossible to obtain the stable products in industrialization.
Visibly, the key obstacle in the free radical-induced polymerization is that there exists always a contradiction between the polymerization rate and the polymerization molecular weight. That is, if to achieve a sufficiently high molecular weight, the polymerization time must extend for too long time, which is unacceptable for the reactive extrusion process in continuous and rapid production. Conversely, the molecular weight will decrease down if the rate of polymerization is to be enhanced. In order to breakthrough this bottleneck, CN103146105B published a technique of reactive extrusion polymerization of (meth)acrylic monomers. Its superior rests with making fully use of auto-acceleration effect in the free radical polymerization process, because only during the auto-acceleration effect interval in the free radical polymerization process, it becomes possible to rapidly enhance both molecular weight and polymerization rate. However, the auto-accelerating effect is a fatal disaster in the usual polymerization kettle and must be avoided as far as possible. Otherwise, the reaction system will immediately generate the burst polymerization, even explosion. However, the twin-screw reactive extrusion is expert in the field of mixing the melt with high viscosity. Because the twin-screw extruder is designed aiming at high-viscosity melt, the extruder can make the melt to be mixed fully, therefore, the heat can be good transfer and the temperature controlled accurately and conveniently. Dispersion and distribution of the mixture are good provided, a small quantity of catalyst and reactants can be mixed equably. Furthermore, because the polymer melt in the screw extruder can get very good surface renewal, which is beneficial to making the small molecule byproducts to be removed out. Visibly, it is the key of the realization of the reactive extrusion polymerization initiated by free radical to fully make use of the autoacceleration effect during the reactive extrusion polymerization. CN103146105B proposed a technique, in which the monomer of acrylic resin, the initiator and the modified resin were fed into the extruder from the first barrel on the first-stage extruder of the double-stage screw extruder. In order to make the polymerization system early enter the aotoacceleration area, the viscosity of the system was changed by regulating the modified resin, thereout, the reactive extrusion polymerization could carry out fleetly. However, there exist still some shortcomings in this technology, such as the ratio of acrylic resin monomer to modified resin is restricted by the viscosity. Second, although the rate of the polymerization may become very high after into the autoaccelerate area, the total quantity of the monomer fed into the extruders limited to the initial added that, more monomers does not be fed into the late barrel of the extruder, therefore, the autoacceleration effect cannot be made good use of, and the production efficiency is restricted severely; Third, the initiator does not be supplementarily added in time after its half-life, delaying the conversion of the polymerization, as a result, productivity also is restricted severely.
Invention Content
The purpose of the present invention is to overcome the above shortcomings of current technique and to provide a technique of the copolymerization, in which both the polymerization rate can fully meet the residence time being very short of the reactive extrusion polymerization, and also meet the request of the high molecular weight for application, which combine the functionalization and high performance in one technique.
The object of the present invention can be achieved by the following technical plan:
A method of the copolymerization of vinyl monomers by the reactive extrusion polymerization, in which the vinyl monomers or at least another vinyl monomer for copolymerization and the initiator are fed to the first barrel on a twin screw extruder and the modified resin is subsequently added in late barrel. The correspond monomers will be supplementarily fed into the barrel after entering the autoacceleration effect area of the system, and the initiator into the barrel corresponding to its decomposition temperature after exceeding its Half-life, as well as the inorganic nanomaterials for modification. The antioxidant and antiultraviolet agent will be added at the end of polymerization, then unpolymerized monomers and byproducts be removed at the devolatilization barrel. The vinyl copolymers with scheduled molecular weight, from 5×102 to 6×105, can be obtained through the reactive extrusion polymerization.
A process of vinyl monomers copolymerization via reactive extrusion is described as following. The vinyl monomers or with at least one kind of vinyl co-monomer and initiator are fed into the first helical segment of the first order twin screw extruder and the modified resin is added into the subsequent screw. Then the corresponding monomers and initiator are fed supplement into the screw segment right after the system enters the auto-acceleration zone, followed by adding micro/nano-inorganic particles then the vinyl pre-copolymer could be extruded and transferred into the second-order twin-screw extruder. The vinyl monomers and initiator are continually added and the polymerization antioxidant and anti-UV agent would be added near the end of the extruder, and the unpolymerized monomers and by-products are removed at the devolatilizing section. The molecular weight of the vinyl copolymer resin could be controlled and adjusted from 5×102 to 6×105 via controlling the temperature of different screw segments during the reaction extrusion polymerization.
The mentioned twin-screw extruder has a twin-screw power inlet with a pressure resistance of 0.3 MPa or more without leakage or a tight-intermeshing twin-screw extruder with a reverse flow structure, which is provided with an inert gas introduction unit and devolatilization section, the inert gas introducing unit introduces the inert gas into the devolatilizing section.
The mentioned first-order twin-screw extruder has a double-screw power inlet with a pressure-resistant pressure of 0.3 MPa or more without leakage, or is a tightly intermeshing twin-screw extruder with a reverse flow structure, the second-order twin-screw extruder is Co-rotating or non-rotating twin-screw extruder; the second-order twin-screw extruder is provided with an inert gas introduction unit and a devolatilization spiral unit, and the inert gas introduction unit introduces the inert gas into the devolatilizing section.
The total mass ratio of the mentioned modified resin to the vinyl monomers and the vinyl comonomers ranges 0-30:100-70; the dosage of the initiator is about 0-20 wt % of the total amount of all the vinyl monomers; The dosage of the mentioned micro/nano inorganic modified material is 0-30 wt % of the total mass of the vinyl monomers; the mass ratio of the antioxidant and the UV-resistant agent is 2:1-1:2. The adding amount of the antioxidants and anti-UV agents is 0.1-1 wt % of the total reactants.
The mentioned vinyl monomers or vinyl comonomers are selected from one or several types of α-methylstyrene, divinylbenzene, acrylonitrile, butadiene, isoprene, methacrylic acid, methylmethacrylate, Ethyl methacrylate, butyl methacrylate, pentyl methacrylate, hydroxyethyl methacrylate, β-hydroxypropyl methacrylate, cyclohexyl methacrylate, glycidyl methacrylate, Acrylic acid, ethyl acrylate, butyl acrylate, pentyl acrylate, hydroxyethyl acrylate, β-hydroxypropyl acrylate, cyclohexyl acrylate, glycidyl acrylate, polycyclic norbornene methacrylate, methacrylic acid Dicyclopentenyl methacrylate, phenyl methacrylate, p-chlorophenyl methacrylate, adamantyl methacrylate, isobornyl methacrylate, vinyl pyridine, maleic anhydride, maleic acid, fumaric acid, Maleic acid monoesters, maleic acid diesters, fumaric acid monoesters, fumaric acid diesters, N-methyl maleimide, N-cyclohexyl maleimide, N-phenyl maleimide, N-tolyl maleimide, N-o-chlorobenzene maleimide, N-Itaconic acid, itaconic acid ester, sorbic acid, sorbic acid ester, tetrafluoroethylene, hexafluoroethylene, vinylidene fluoride, vinyl chloride, vinylidene chloride, vinyl isocyanate or acryloyl chloride.
Among these vinyl monomers, N-phenylmaleimide, dicyclopentenyl methacrylate, phenyl methacrylate, p-chlorophenyl methacrylate, adamantyl methacrylate and the like are copolymerized, can significantly improve the polymer's glass transition temperature and modulus, resulting high performance products. Meanwhile, methacrylic acid, acrylic acid, maleic anhydride, maleic acid, fumaric acid, glycidyl methacrylate, etc. involved after the copolymerization, could endow various functions to the copolymers. The content and distribution of these co-monomers in the macromolecular chains are determined by their reactivity ratio r1, r2, r3 If all the reactivity ratios are far less than 1, then their content and distribution could directly controlled via adjusting the monomer addition. However, if the reactivity ratio of a certain monomer is large or even much larger than 1, then the amount of the monomer should be added to the barrel of the back stage instead of the previous stage of the extruder. Thus, uniform distribution of the monomers in the final copolymers could be achieved.
The content and distribution of each component in the vinyl copolymer can be measured by the commonly used infrared spectroscopy and nuclear magnetic resonance spectroscopy according to the structure of each monomer composition.
The initiators are selected one or several from inorganic peroxide, organic peroxide, azo initiators or redox initiators;
The mentioned inorganic peroxides include potassium persulfate, sodium persulfate and ammonium persulfate;
The mentioned organic peroxide has the general formula: R—O—O—R′, wherein R and R′ are H, alkyl, acyl or carbonate, R and R′ can be the same or different;
The mentioned azo initiators include azobisisobutyronitrile and azobisisobutyronitrile;
The mentioned redox initiators include cumene hydroperoxide-ferrous salt and organic peroxide-tertiary tertiary amine system.
Among the above initiators, the preferred are those which can be dissolved in the monomers, or at least to be able to dissolve in some of the solvents, which can facilitate the quantitative addition and form a homogeneous polymerization system in the twin extruder. In addition, with the increase of polymerization conversion, the viscosity of the system is also constantly increasing, it is necessary to accordingly increase the barrel temperature. Therefore, a single initiator cannot address the requirement of constant active species concentration during the polymerization. The initiator added before the extruder requires a lower decomposition temperature, and after the half-life is reached, the decomposition temperature of the later-added initiator should be increased with the increase of the barrel temperature. Therefore, the sequential supplement of the initiators, makes the polymerization under constant concentration of active species. The copolymer obtained is more uniform and stable.
The mentioned modified resins are selected one or several polymers from butadiene styrene rubber, nitrile rubber, natural rubber, styrene-butadiene-styrene triblock copolymer, styrene-isoprene-styrene triblock copolymer, hydrogenated styrene-butadiene-styrene triblock copolymer, hydrogenated styrene-isoprene-styrene triblock copolymer, styrene/butadiene random copolymer-based thermoplastic elastomer, methyl methacrylate-butyl methacrylate micro nanoscale diblock copolymer, methyl acrylate-butyl acrylate micro-nano-level diblock copolymer, methyl methacrylate-butyl acrylate micro-nano-level diblock copolymer, methyl methacrylate/butyl methacrylate random copolymer, methyl methacrylate/butyl acrylate random copolymer, styrene-butadiene-methyl methacrylate micro-nano-level triblock copolymer, styrene-isoprene-methyl methacrylate micro-nano-level triblock copolymer, thermoplastic polyurethane, and homopolymer or copolymer of a vinyl monomer and comonomers. The addition of the modified resin to the reactive extrusion polymerization system can increase the viscosity of the system, accelerate the emergence of auto-acceleration effect and accelerate the polymerization reaction. Moreover, if the modified resins are rubber or thermoplastic elastomer, the impact resistance of the copolymer will be significantly improved. If the block copolymer is added to the polymerization system, the copolymer can also be used as an excellent polymer compatibilizer.
The mentioned nano-micron inorganic particles are selected from one or several kinds of micro-nano silica, micro-nano calcium carbonate, micro-nano aluminum oxide, micro-nano aluminum hydroxide, micro-nano tantalum pentoxide, micro-nano whiskers, micro-nano quartz, micro-nano cerium oxide, micro-nano europium oxide, micro-nano zirconia, micro-nano barium oxide and micro-nano lanthanum oxide. Addition of a very small amount of nano-micron inorganic particles can not only increase the rigidity, but also improve the toughness of the copolymer to a certain extent. Moreover, the addition of micro-nano whiskers or micro-nano zirconia, can improve the scratch resistance of the copolymer surface; the addition of micro-nano cerium oxide, micro-nano europium oxide, micro-nano barium oxide or micro-nano lanthanum oxide can increase the ability of the copolymer to resist thermal oxidative degradation and obtain functions of fluorescence, light scattering, light-emitting, piezoelectric and the like.
The mentioned antioxidants are selected one or several kinds from commercially available antioxidant 168, antioxidant 1076, antioxidant bht, antioxidant B215, antioxidant 245, antioxidant 1010, thiodipropionate dioctadecyl ester, diphenyl isooctyl phosphite, tetrakis[methyl-β-(3,5-di-tert-butyl-4-hydroxyphenyl) propionate] pentaerythritol ester, 1,3-tris(2-methyl-4-hydroxy-5-t-butylphenyl) butane, 1,1,3-tris (2-methyl-4-hydroxy-5-t-butylphenol), 3,5-di-tert-butyl-4-hydroxyphenyl) propionate, 2,2′-methylenebis (4-methyl-6-tert-butyl) phenol, 4′4-thiobis (6-t-butyl o-cresol), 4,4′-thiobis(3-methyl-6-t-butyl)phenol, 4′4-(dihydroxy-3,3′,5,5′-tetra-t-butylbiphenyl). The anti-UV agents are selected one or several kinds of UV-531, UV-9, UV-326, UV-327, UV-328 and UV-329, phenyl hydroxybenzoate, o-nitroaniline or p-cresol.
The mentioned twin-screw extruder has an aspect ratio over than 48/1 and a screw speed of 2-300 rpm.
The mentioned inert gas includes carbon dioxide, nitrogen, helium, argon, butanol, alcohol or water vapor. The choice of inert gas is very important, it is based on the principle that the polarity of inert gas should be as close as possible to the polymer, so as to more efficiently remove the unpolymerized monomers and by-products from the products and avoid as far as possible the inertness with phase transition gas. Water is often used as an auxiliary gas for removal byproducts. The boiling point of water is usually 100° C. at the atmospheric pressure, and the temperature of the extruder is about 200° C. Although it is convenient to be realized in the equipment, the vaporization of water will cause the great energy consumption and solidify the polymers, resulting more difficult to remove by-products. Therefore, carbon dioxide should be the most preferred.
The mentioned reactivity ratio between the vinyl monomer and the vinyl comonomer should be less than or equal to 1, otherwise, monomers with lower reactivity must be added into different screw segments according to the content requirement in the copolymer.
The absolute molecular weight of the above vinyl copolymer obtained via reaction extrusion polymerization can be measured by gel permeation chromatography (GPC).
In order to determine the absolute molecular weight and its distribution of the polymer, multiple-detection system with multiple angle laser detection (LS) and refractive index differential scanning (RI) must be employed, i.e. Wyatt Technologies Water 1515 multi-detection gel permeation chromatography (United States). Tetrahydrofuran (THF), dimethylformamide (DMF), dimethylsulfoxide and the like can be used as the mobile phase at a flow rate of 1.0 ml/min and a detection temperature of 25° C. The column was filled with styrene-divinylbenzene copolymer gel. The mass concentration of the sample solution is 0.002˜0.004 g/ml.
The mentioned structural principle of the co-rotating tightly-intermeshing twin-screw extruder with the reverse flow structure is shown in FIG. 1. It is actually a twin-screw pull-out machine. Therefore, the thrust bearing in the gearbox has the opposite bearing direction compared to the normal one. The raw materials are added into the extruder through the screw tip, and the output material is close to one end of the gearbox. The aspect ratio is 48/1 or more, preferably 60/1 or more, and the screw speed is 2-300 rpm. The temperature of each screw segment should be successively increased to the last stage of the extruder, the temperature ranges 100-300° C., higher than the flow temperature of the copolymer.
Twin-screw extruders have a strong ability to transfer viscous melts, but poor capacity for transferring water-like monomers. Therefore, in order to ensure long-term stable operation, the extruder pressure at the entrance must be above 0.3 MPa. The using of a pressure metering pump can not only pump the monomers into the screw, but also promote the monomer with the screw rotation forward.
The extruder using the reverse direction structure is completely closed and there is no leakage problem because the monomers are entered into the extruder at the end of the screw. The output port has turned to one end close to the gearbox. There are two screws running, but at that time, the material has become a high viscous melt, and there is almost no leakage problem.
The inert gas introduction unit of the present invention is shown in FIGS. 2 and 3, whereas 1 is a vacuum depot, 2 is an inert gas entry, and 3 is a unidirectional needle valve. By introducing the inert gas to reduce the partial pressure and has the effect of carrying, it can efficiently remove volatilization components.